In-line hydrotreatment process for low TAN synthetic crude oil production from oil sand

ABSTRACT

A low cost and energy efficient process for manufacturing low TAN synthetic crude oils from oil sands, by selective removal of organic acids from carboneous distillates and or blends of carboneous distillates in a hydroprocessing unit integrated within the process flows of a bitumen Upgrader.

FIELD OF THE INVENTION

The invention relates to processing bitumen to reduce acid content.

BACKGROUND OF THE INVENTION

Organic acids (Naphthenic, aromatic and paraffinic carboxylic acids) in crude oils have been demonstrated to strongly influence the corrosion rate in refining equipment. The acidity of a crude oil is typically measured as the Total Acid Number (TAN) by ASTM Method D664 or the UOP 565 Procedure. It is strongly desirable to reduce the TAN of the crude oil as early in the oil refining process as possible, to minimize corrosion impact on the integrity of the refining equipment.

It is well known that crude oils, and crude oil fractions contain sulfur, nitrogen, and other compounds, and a large number of processes have been proposed for the removal of such compounds from crude oils and or crude oil fractions. However, no single process or solution is useful, workable and economic for all crude oils as the organic components of crude oils vary depending on the source of the hydrocarbon. Consequently, processing methods and equipment must be arranged to be effective for the crude oil constituents that are processed. While the organic chemical composition of fuel oil and lubricating oil products is relatively uniform as supplied to the end use consumer, the hydrocarbon start materials have diverse constituents that vary with the many locations on the earth that the hydrocarbon start materials are recovered from.

A particularly challenging hydrocarbon start material is bitumen as it presents numerous difficulties and virgin bituminous fractions have organic characteristics that are quite unique to bitumen as a source of oil products. Virgin Bitumen distillates are produced by vacuum distillation processing of bitumen. Typically, virgin bituminous distillates are quite different from distillates obtained from conventional crude oil sources. Virgin bituminous distillates contain very high concentrations of various ringed molecular structures, and are very low in hydrogen content. For example, Canadian Athabasca tar sand bitumen usually contains about 95% ringed molecular structures as compared to a 10-50% ringed molecular structure content found in conventional crude oil hydrocarbons. Moreover, distillates or fractions derived from bitumen are further characterized by a very high molecular weight and by having a high density, high viscosity, low viscosity-index and low fluidity properties. These unique properties of bitumen derived virgin fractions, particularly under lower hydroprocessing severity, negatively affect the reactor hydrodynamics, resulting in lower mass transfer rates; and hence render them more difficult to upgrade into synthetic crude oil.

Heretofore it was not known if organic acids in bitumen derived virgin distillates could be selectively removed at hydroprocessing severity conditions that are below the conditions which result in the onset of sulfur and other reactions.

Moreover, the organic characteristics of bitumen hydrocarbons themselves vary from one location to another. Bitumen located in Canada has chemical properties and characteristics that are different again from the organic characteristics of bitumen from other known sources in the world. Right from the outset, recovery of relatively small amounts of bitumen from tar sands sources in Canada requires handling vast amounts of sand and separating the bitumen from the sand. Once the bitumen is separated from the sand, the bitumen must then be upgraded into a synthetic crude oil to enable production of oil products from the bitumen.

In the processing of bitumen, bitumen fractions are produced, which are very high in organic acids. Sour Synthetic crude oils blended from these fractions are very high in TAN and also contain very high concentrations of sulfur, nitrogen and other undesirable compounds. Severe hydrotreatment of the virgin bitumen fractions, independently or in blends with other fractions produced from thermo or other conversion processes, is necessary to remove the undesirable compounds and reduce the TAN such that sweet, low TAN, synthetic crudes can be blended for sale to refineries for conversion into fuel products.

One of the traditional approaches to reducing the acid content that is used in processing facilities of other types of crude oils is to use chemical neutralization, where various bases are added to the crude oil to neutralize the acidic components. Unfortunately, however, this approach has not been found to be successful when applied to bitumen processing. In the context of bitumen processing, this approach introduces other processing problems such as emulsion formation, increases in the organic salts, particularly those of calcium, magnesium and sodium, which further exacerbates corrosion and conversion issues in down-stream upgrading and refining process units.

Another approach is to refine crude oils into products even though the crude oils contain high TAN components. In this approach, corrosion-resistant metals are used in the construction of refining units, which results in specialized refining facilities each with significant increased capital investment to provide the corrosion-resistant units. Moreover, this approach is prohibitively expensive to retrofit onto existing refining facilities due to changes in component parts, increased component costs, changes in process flows and changeover production losses. Consequently, this approach is not in widespread use.

Another approach is to add corrosion inhibitors to the crude oil to protect the metallurgy of the refining units, which often results in other processing complications in down-stream units such as catalyst poisoning and/or inhibition, or fouling in furnace tubes and other equipment etc.

Yet another approach is to blend high TAN crude oils with lower TAN crude oils to reduce the TAN of the output crude oil and manage the corrosion rate at an acceptable level in that manner. This approach results in high inventory costs and greatly increases logistical and feed supply costs, for example sourcing and obtaining delivery of lower TAN crude oils for blending.

The following patent documents relate to one or more of these approaches in dealing with TAN components of conventional crude oils.

U.S. Pat. No. 5, 985,137 describes a process to upgrade conventional crude oils by destruction of naphthenic acids, removal of sulfur and removal of salt by mixing with alkaline earth metal oxides to convert substantially all of the naphthenic acids contaminants to no-acid compounds, and alkaline earth metal carbonates, and also to convert the sulfur contaminants to alkaline earth metal sulfide.

U.S. Pat. No. 6,531,055 B1 describes a process for extracting naphthenic acids using solvent systems comprising liquid alkanols, water and ammonia, to facilitate selective extraction and easy separation from conventional whole crude.

U.S. Pat. No 5,961,821 describes a process for extracting organic acids, heavy metals and sulfur from a starting crude oil comprising of treating the crude oil with ethoxylated amine and water under specified conditions and residence time to form water-in-oil emulsion of amine salt for separation from the treated crude oil.

Similarly, U.S. Pat. No. 6,096,196 describes the same process as U.S. Pat. No. 5,961,821 by treating the crude oil with alkoxylated amine and water.

U.S. Pat. No. 4,634,519 similarly describes a process for extracting naphthenic acids using solvent systems comprising liquid alkanols, water and ammonia, to facilitate selective extraction and easy separation from crude oil fractions prepared by distillation.

The following patent documents relate to various hydroprocessing processes that have been proposed to reduce crude oil TAN.

U.S. Pat. No. 2,921,023 is directed toward a method of improve catalyst activity maintenance during mild hydrotreating to remove naphthenic acids in high boiling petroleum fractions. The catalyst is molybdenum on a silica/alumina support wherein the feeds are heavy petroleum fractions.

U.S. Pat. No. 2,734,019 describes a process for treating a naphthenic lubricating oil fraction by contacting with a cobalt molybdate on a silica-free alumina catalyst in the presence of hydrogen to reduce the concentration of sulfur, nitrogen, and naphthenic acids.

U.S. Pat. No. 3,976,532 relates to a very mild hydrotreatment of virgin middle distillates from crude oils in order to reduce the total acid number or the mercaptan content of the distillate without greatly reducing the total sulfur content using a catalyst which has been previously deactivated in a more severe hydrotreating process.

Selective hydrogenation to remove essentially the naphthenic acids without hydrogenating other compounds from crude oil and portions of crude oil has been reported in CA. U.S. Pat. No. 2,198,623 and also in U.S. Pat. No. 6,063,266. U.S. Pat. No. 5,910,242 describes a process for reducing the Total Acid Number in crude oil by hydroprocessing the crude oil in front of a refinery crude tower. These hydroprocessing processes reduce the corrosivity of crude oils by hydrotreatment in various process arrangements. The process arrangements include stand-alone hydrotreaters dedicated to naphthenic acid removal, or a refinery facility having a naphthenic acid hydrotreater placed in the crude oil processing process flow before the refinery crude tower. Thus these hydroprocessing processes are arranged in a refinery environment to provide equipment to process high TAN crude oils while reducing or eliminating corrosion caused by the high TAN crude Oil.

It is well known in the oil sand industry, where bitumen is first extracted from tar sands and then upgraded to synthetic crude oils, that virgin bituminous fractions are very high in organic acids. Consequently, sour synthetic crude oils blended from these virgin bituminous fractions are very high in TAN and also contain very high concentrations of sulfur, nitrogen and other undesirable compounds. Severe hydrotreatment of the virgin fractions, independently or in blends with fractions produced from thermo or other conversion processes, removes the undesirable compounds and reduces the TAN such that sweet, low TAN, synthetic crudes can be blended for sale to refineries for conversion into fuel products.

Severe hydrotreatment to upgrade bituminous derived fractions is a capital-intensive process. In addition to the hydrotreater, hydrogen production and sulfur recovery units are also required. Hence it is strongly desired for the oil sand industry to be able to produce a low TAN synthetic crude oil where particularly organic (naphthenic, aromatics, and paraffinic) acids from virgin bitumen derived distillates and blends of such virgin bituminous distillates, can be removed economically.

SUMMARY OF THE INVENTION

The present invention provides a low cost process for the removal of the organic (naphthenic, aromatic, paraffinic carboxylic) acids during the synthetic crude oil manufacturing process. In accordance with the invention, low TAN synthetic crude oils can be produced from bitumen start material, which then eliminates the corrosion concerns of the refining industry when synthetic crude oils are refined into fuel products in refineries.

It has now been discovered, that virgin low TAN synthetic crude oils can be produced by integrating a selective organic acid (naphthenic, aromatic and paraffinic carboxylic acids) hydroprocessing unit within the synthetic crude oil manufacturing process. By providing a selective organic acid hydroprocessing unit, other hydroprocessing conversion reactions for sulfur and other undesirable compounds that occur in a severe hydroprocessing unit are not necessary and can be kept to a minimum.

It is possible to carry out selective removal of the organic acids from the vacuum fractionated cuts of virgin bitumen fractions by selective hydrogenation of the organic acids under very mild conditions. Under such mild conditions, any substantial amount of desulfurization reactions or denitrification reactions or saturation reactions is avoided, which results in a moderate hydrogen consumption. It has further been discovered that a comparative low hydrogen purity (>50%) in the hydrotreating gas will effect good conversions. As a consequence, common hydrotreater bleed gases may be used for the hydrogenation process thereby eliminating the need for hydrogen production units or equipment.

In one of its aspects the invention provides the selective hydrogenation of the organic acids under very mild conditions using a low purity hydrotreating gas. The low purity hydrotreating gas is sourced from waste gases of the bitumen processing plant into which the selective hydrogenation equipment is integrated with. Integration of the selective hydrogenation equipment with a bitumen processing plant achieves numerous advantages resulting in lower costs are achieved relative to the common art of using high purity hydrogen as a hydrotreating gas.

The present invention provides a process for removing naphthenic and other carboxylic acids from bitumen derived distillates and blends of such distillates. In accordance with the invention, the facilities to carry out the process are incorporated into the bitumen processing facility that upgrades the bitumen into synthetic crude oil blending components. Incorporation of the process facilities into a bitumen processing facility eliminates the need for separate hydrogen production and process heat supply/cooling inputs to carry out the selective hydrotreatment process. Thus reduced TAN of bitumen derived vacuum distillates and blends of bitumen derived vacuum distillates is achieved. The process of the invention can be integrated into the facilities of a bitumen Upgrader that receive feed bitumen or diluted bitumen streams that are produced from bitumen sand excavation and Clarke hot water process or other bitumen extraction or processes from Steam Assisted Gravity Drainage (SAGD) or other bitumen production methods.

If the bitumen is produced by excavating the tar sands formation material and then extracting it from the sand using a caustic hot water process (i.e. Clarke hot water process) or other organic/inorganic solvent process, a pre-treatment of the bitumen derived distillates or blends of bitumen derived distillates to demulsify, dewater, and demineralize the bitumen feed material is generally necessary prior to distillation. However, if the bitumen is produced from tar sands by Steam Assisted Gravity Drainage (SAGD) or other in situ thermally assisted gravity drainage bitumen production methods, it may not be necessary to process the bitumen to demulsify, dewater, and demineralize it prior to distillation. Irrespective of the manner in which the bitumen was produced, it is generally advantageous to dilute the feed bitumen with hydrocarbon solvent or other diluent.

The process of the invention achieves a selective reduction of the content of organic acids in bitumen derived distillates, or blends of bitumen derived distillates to less than about 0.45 mg KOH/g without the simultaneous hydrogenation of sulphur compounds and nitrogen compounds which may be present.

In accordance with the invention, apparatus to carry out the process is integrated within a bitumen Upgrader that processes bitumen recovered from tar sand into synthetic crude oils. Thus, the invention provides a low capital and operating cost solution by integrating the process within a bitumen Upgrader, and in particular integrating the process with a bitumen vacuum fractionation unit of the bitumen Upgrader. In the preferred arrangement, the hot vacuum gas oils produced by the bitumen vacuum fractionating unit are diverted from their run-down heat exchangers and supplied directly as a hot feed to an in-line hydrotreating reactor. Thus, process heaters to heat the hot feed of the hydrotreating unit are eliminated. The in-line hydrotreating reactor product is then supplied to the run-down heat exchangers of the bitumen fractionating unit vacuum tower to cool down the hydrotreated vacuum gas oils. Thus, separate process heat exchangers to cool the product of a hydrotreating unit are eliminated. This configuration of the in-line hydrotreating reactor arranged with the fractionating unit vacuum tower allows the vacuum tower product to be heated and cooled without additional equipment by using the exiting process facilities of the bitumen fractionating unit vacuum tower.

The novel integrated configuration design provides a low capital and operating cost solution that achieves reduction of the acidity of bitumen distillates to obtain low TAN synthetic crude oils. The process is incorporated in a bitumen upgrading facility thereby eliminating the need for TAN treatment in downstream refineries that process the synthetic crude oil products obtained from the bitumen start material.

Thus, the invention provides a process for the manufacturing of low TAN synthetic crude oils from oil sand derived bitumen streams by hydrotreatment in a selective hydro-deoxygenation processing facility nested within a bitumen vacuum distillation unit. A preferred embodiment of the invention will now be described with reference to the attached drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1—Is an overview process flow schematic diagram of a prior art bitumen processing plant drawn to identify process flows and equipment that accommodate apparatus adapted to carry out the process of the invention.

FIG. 2—Is a process flow schematic diagram of the preferred embodiment of an arrangement of apparatus adapted to carry out the process of the invention including a conventional bitumen vacuum fractionation unit and an inline hydro-deoxygenation reactor unit interoperably connected to it.

FIG. 3—Is a process flow schematic diagram of an alternate embodiment of the inline hydro-deoxygenation reactor unit of FIG. 2 interconnected with a conventional bitumen fractionation unit.

FIG. 4—Is a graph of TAN of bitumen distillates after processing by experimental pilot plant equipment arranged to carry out the process of the invention operating under various conditions.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT

FIG. 1 shows an overview process flow schematic diagram of a bitumen Upgrader bitumen processing facility drawn to show the apparatus and process flows that are used to interconnect with to carry out the process of the invention. A geographic formation 10 includes a bitumen containing tar sand formation that has a source of bitumen. The bitumen is produced in one manner by mining or excavation and is then transported to and processed in a bitumen upgrading processing facility. When the bitumen is produced by excavation, the bitumen bearing tar sand is removed from the earth at 12 and transported by truck 14 to a receiving facility. One form of receiving facility is a slurry transport pipe 16, which transports the bitumen sand to the intake facilities of the bitumen Upgrader where the bitumen is separated from the sand in primary separation cells 18. In the primary separation cells 18 the received bitumen and sand material is mixed with a caustic hot water. The bitumen is separated out of the slurry solution using air flotation producing a bitumen froth output 20.

The bitumen froth output 20 from the primary separation cells is supplied to a secondary separation facility 22, which removes the water and mineral fines present in the bitumen froth to obtain an enhanced bitumen product. During processing in the secondary separation facility 22, a diluent 24 is also added to assist in the separation process by reducing the viscosity of the bitumen and enhancing the purification of the bitumen produced by the secondary separation facility 22. The bitumen product from the secondary separation facility is processed by a diluent recovery unit 26 to remove the diluent 24 that was added to the bitumen in the secondary separation facility to assist in the separation process. The diluent recovered by the diluent recovery unit 26 is then recycled back to the secondary separation facility as part of the diluent supply 24. The output of the diluent recovery unit is a bitumen feed 28.

An alternate method of bitumen production includes drilling a well 46 to supply steam to the bitumen sand formation from a source of steam 48. The steam heats the bitumen in situ and increases its flowability causing it to pool in the lower portion of the volume treated by the steam. This form of bitumen production is referred to as the steam assisted gravity drainage method (SAGD). The bitumen that pools in the formation is extracted from the well and is then transported by pipe 50 directly to the diluent recovery unit 26 of the bitumen Upgrader. Unlike the bitumen produced using the mining and flotation method, the SAGD produced bitumen does not require separation from enormous quantities of sand by a primary separation cell facility. As with bitumen that is extracted using mining and flotation methods 10, 12, 14, 16, 18, 20, the SAGD recovered bitumen is processed by providing a diluent at 48 to reduce the viscosity and increase the flowability of the recovered bitumen.

Irrespective of the method used to produce the bitumen, once the bitumen has been processed in the DRU 26, the bitumen feed 28 is supplied to a vacuum distillation unit 52, where virgin bitumen distillates (such as virgin kerosens, diesels, and gas oil cuts) 54 are produced. These virgin bitumen distillates have a high TAN content and have heretofore been used in blending sour synthetic crude oil.

The residual bitumen 55 is then fed to a thermal conversion unit 30, for example a coker unit, for conversion to lighter hydrocarbons. The products from the thermal conversion or coker unit 30 are separated into coker products streams based on boiling points ranges. From the coker unit, a coker gases stream 32 is produced as well as other streams of compounds with differing ranges of boiling points, including a stream 34 of naphtha compounds that have a boiling point less than about 315° F. Also produced are a stream 36 of diesel compounds which have higher boiling points in the 300-650° F. range and a stream 38 of gas oils that have boiling points in excess of 600° F.

The distillate compounds produced by the coker unit 30 are low in hydrogen content and are high in Sulfur, Nitrogen and other undesirable constituents. Consequently each of these compounds is further treated in the bitumen upgrading process by supplying each of the compounds streams to a corresponding hydro-treatment facility 40. After the hydro-treatment process, the hydrotreated petroleum products streams, namely, the paraffinic gas oils stream PG, the paraffinic diesels stream PD and the paraffinic naphtha stream PN, are supplied to a corresponding storage facility 42. For product delivery, the petroleum products are drawn from storage and are blended at blending facility 44 to produce sweet low TAN synthetic crude for supply to downstream refiners.

The invention provides a bitumen processing facility that includes a hydro-deoxygenating facility 74 which provides a mild hydro-treatment to effect TAN reduction in a manner that will be explained in more detail with reference to FIGS. 2 and 3. The process of the invention produces low TAN gas oil products for sour low TAN synthetic crude oil blending which overcome the difficulties that are present when high TAN gas oil products are processed in downstream refinery processing.

Thus the process of the invention is carried out in the facilities of a bitumen Upgrader to remove essentially organic acids from the virgin distillate hydrocarbon oils derived from bitumen or diluted bitumen. In accordance with the process of the invention, virgin bitumen distillates, or blends of such distillates, are separated from the bitumen feed of the bitumen Upgrader and then hydrogenated at very mild temperature over a catalyst. The catalyst is of a kind used for hydrogenation of vacuum gas oils and/or atmospheric residue, and preferably is a catalyst consisting of nickel-molybdenum or cobalt-molybdenum, deposited on alumina as a carrier material. The process is carried out by a bitumen vacuum fractionating unit that interoperates with the facilities of a bitumen Upgrader producing:

-   -   (a) a distillate derived directly from (Althabasca) bitumen; or     -   (b) a blend of distillates, in any ratio which has previously         been distilled into fractions, from (Althabasca) bitumen

In the process of the invention it is preferred to carry out the hydrogenation at a pressure range of 300-650 PSIG, at a temperature range of 350-600° F. preferably in the range of 400-530 degrees Fahrenheit, at a Liquid Hourly Space Velocity (LHSV) range of 0.1-5.0, being the ratio of the volume of feed divided by the volume of catalyst, and with a charge gas supply rate range of 250-1500 standard cubic feet per barrel (SCFB).

The hydrogenation is suitably effected in one or more parallel reactors or one or more reactors arranged in series, each reactor having one or more fixed catalyst beds. As mentioned, the catalysts utilized in the process of the invention are such catalysts that have proved to be suitable for hydrogenation of gas oils and atmospheric residue oils. To carryout the mild hydrogenation process in a bitumen processing facility successfully, it is important that the carrier material of the catalyst is sufficiently porous to allow penetration by diffusion of even the heaviest part of the bitumen derived distillates or blends of bitumen derived distillates into the catalyst pores. Therefore, the carrier material should have porosity such that the final supported catalyst preferably has a porosity of the magnitude 10 to 12 nanometers (nm). Particularly useful catalysts comprise nickel-molybdenum or cobalt-molybdenum deposited on alumina as a carrier material. The bitumen derived distillate or blends of bitumen derived distillates flow rate through the catalyst is preferably 0.5 to 5.0 LHSV and most preferred 1.0 to 3.0 LHSV.

On exit from a vacuum faction distillation tower unit 52, the hot virgin distillates while at distillation temperature are supplied to a hydrogenation reactor. In the preferred embodiment, the distillates are supplied directly to a hydrogenation reactor along with a hydrogen rich (>50% H2) gas for processing at the conditions just specified. In an alternate embodiment, the hot distillates are sent to a surge tank. Distillates drawn from the surge tank are pumped to processing pressure and then mixed with hydrotreater bleed gas containing at least 50% H2 at a rate in the range of 500-1500 SCFB. The mixture is supplied to the hydrogenation reactor for processing directly. Depending on the temperature of the hot virgin distillate(s), and the molecular species of organic acids as well as the properties of the virgin distillates, booster heater may be added for higher reaction temperature if required.

An example of a facility arranged in a preferred manner to embody the process flows of the invention is described in more detail herein. The main features of the preferred embodiments are shown in FIGS. 2 and 3.

In FIG. 2 a bitumen feed 28 is supplied to a vacuum fractionator vacuum distillation tower unit 52. The vacuum fractionator distills the bitumen feed into a hot virgin Light Vacuum Gas Oil (LVGO) stream 58 and Heavy Vacuum Gas Oil (HVGO) stream 60, at 365 degrees Fahrenheit and 520 degrees Fahrenheit respectively. The high TAN LVGO and HVGO streams are taken off the vacuum distillation tower unit 52 by pumps 62 and 64. In the preferred embodiment of FIG. 2 tap points 66, 68 are provided in the LVGO and HVGO process streams and portions 70, 72 of the LVGO, HVGO streams are taken off the output process streams of the vacuum distillation tower unit 52. The portions 70, 72 of the LVGO and HVGO streams that are taken, are taken either alone or by combining the stream portions together in various volume ratios. The amounts taken and any combining effected varies and is determined by what is found to be useful to obtain optimal processing of the time-varying constituents found in the bitumen feed 28 and the portions taken can include all of LVGO and HVGO streams in their entirety. The portions taken of the LVGO and HVGO streams 70, 72 are supplied to an in-line hydro-deoxygenation hydroprocessing unit 74.

The portions taken off the distillate streams from the vacuum distillation tower unit 52, or blends of the distillate streams, while still hot from egress from the vacuum distillation tower unit 52, are pressured up to hydrogenation pressure range of 500-650 PSIG by a charge pump system 76. A hydro deoxygenation heater (HTR) 78 may be provided if desired depending on the process needs such as: target TAN level, feed-quality, catalyst consumption/aging rate.

The pressurized distillate or distillate blend is mixed with a hydrogen charge gas 80, which is obtained from a source of hydrogen gas. In the preferred embodiment, the source of hydrogen gas is the pressurized waste gas from other hydroprocessing units found in a bitumen Upgrader that the hydrotreater process system is deployed in. The hydrogen charge gas preferably contains at least 50% hydrogen and is supplied at a rate in the range of 100-1000 standard cubic feet per barrel (SCFB), and preferably at a rate in the range of 400-700 SCFB. The actual supply rate will vary depending on the hydrogen content of the charge gas and other operating parameters. Where the source of the charge gas 80 is obtained from other hydroprocessing units of the bitumen Upgrader, it preferably contains 5-6% of H2S to enhance the reactions of organic acid conversion within the in-line hydro-deoxygenation reactor unit 74.

The mixture of the vacuum gas oil liquids 70, 72 and hydrogen charge gas 80 is supplied to an in-line hydro-deoxygenation reactor unit 74. Depending on the temperature of the hot virgin distillate(s), the molecular species of organic acids, as well as the properties of the virgin distillates and target TAN content, a booster heater (HTR) 78 may be added for higher reaction temperature if required. The hydro-deoxygenating reactor unit 74 has a catalyst bed(s) of sufficient size and is loaded with catalyst(s) proven to remove organic acids from the feed. The treated vacuum gas oil (VGO) reactor effluent 82 exits the reactor to a gas-liquid separator 84. The liquid output 86 of the gas-liquid separator is returned to a return tap point 88 to supply the reduced TAN liquid output 86 to an output stream of the vacuum distillation tower unit where it will continue in the downstream process of that system. The heat in the product fluids is recovered by heat exchangers 90. The heat recovered is typically then supplied or recycled to heat at 91 the bitumen feed 28 of the vacuum distillation tower unit 52.

The waste gas 92 of the gas-liquid separator is returned to the originating bleed gas treatment system that it was supplied from at 80. The bleed gas treatment system is present in the facilities of a bitumen Upgrader and provides treatment of bleed gas by hydrogen recovery and/or sweetening for fuel gas production.

The cooled product liquid continues through to the bitumen Upgrader vacuum distillation tower unit rundown system 94, which provides additional cooling, or recycling to the vacuum distillation tower, as needed prior to reporting to tankage 96. The lower organic acid product collected in tankage 96, as measured by Total Acid Number (TAN), is used to blend sour low-TAN crude for transport to market.

If hydrogen rich waste gas 80 from another hydrotreater is not available, or in insufficient quantity for once through operation, or to maximize utilization of available hydrogen in the waste gas, a recycle gas circuit complete with a compressor may be employed to recover the hydrogen containing gas from the hydro-deoxygenation reactor effluent.

Suitable process equipment and suitable safe operating procedures for carrying out the process of the invention as described herein is available from suppliers of the equipment utilized in well-known processes for hydrogenation of gas oils. It is to be noted, however, that additional equipment, which is used in connection with gas sweetening, sulphur recovery and nitrogen removal, is not contemplated or required to carry out the process of the invention.

FIG. 3 shows an alternate embodiment of an in-line hydro-deoxygenation unit. In the embodiment of FIG. 3, the portions of the LVGO and HVGO streams 70, 72 taken off the vacuum distillation unit are supplied to a feed drum 98. The portions of the LVGO and HVGO streams that are taken, are taken either alone or by combining the streams together in various volume ratios. The amounts taken and any combining effected varies and is determined by what is found to be useful to obtain optimal processing of the time-varying constituents found in the bitumen feed 28. A temperature and flow controller 100 is preferably provided to control the flow rates of the portions 70 and 72 of the LVGO and HVGO streams that are supplied to the in-line hydro-deoxygenating unit 74.

The portions 70, 72 taken off the distillate streams from the vacuum distillation unit 52, or blends of the distillate streams, while still hot from egress from the vacuum distillation unit 52, are pressured up to hydrogenation pressure by a charge pump system. In the embodiment of FIG. 3, the charge pump system has a charge pump 102 disposed at the outlet of the feed drum 98. The pressurized distillate or distillate blend is mixed with a hydrogen charge gas 80 and the mixture is supplied to the in-line hydro-de-oxygenation reactor unit 74. The treated vacuum gas oil (VGO) reactor effluent 82 exits the reactor to a gas-liquid separator 84. The liquid output 86 of the gas-liquid separator is supplied to the return tap point 88 where it is incorporated into an output stream of the vacuum distillation tower to continue in the downstream process of that system.

Apparatus to carry out the process of the invention provides a low severity hydrogenation, or hydro-deoxygenation, unit that is integrated in operation with a bitumen processing facility. The hydrogenation unit is placed within the process flows of a bitumen processing facility to obtain operating efficiency and reduced processing cost in processing the bitumen into synthetic crude oil components. Operating efficiency and reduced processing cost is achieved through several benefits obtained by integration with a bitumen Upgrader. A significant capital and operating cost saving is achieved by obtaining the hydrogenation process hot supply feed at operating temperature from the process flows within the bitumen Upgrader, which eliminates the need for a separate feed or charge heater. Consequently the fuel consumption for the hydrogenation/hydro-deoxygenation reactor is eliminated. In certain configurations, such as that shown in FIG. 3, where the supply feed may cool during residency in a feed drum, a charge heater may advantageously be provided. A charge heater may also be used to advantage where the cost of providing and operating a charge heater to obtain elevated process temperatures is offset by a cost reduction obtained by a reduction in size of the hydro-deoxygenation reactor needed to operate at the higher temperature.

Other operating efficiency and reduced processing cost reductions achieved through integration with a bitumen Upgrader include.

-   -   The hydrogenation reactor product cooling is integrated with the         distillation product cooling system for better process cooling         and heating efficiency in product rundown to tankage, resulting         in significant reductions in capital cost and operating costs,         including savings in lower maintenance requirements.     -   A hydrotreater charge gas containing a comparatively low         hydrogen content is advantageously used, thereby reducing or         eliminating the need for stand-alone or additional hydrogen         production and or purification facilities to support the low         severity hydrogenation TAN reduction process in the bitumen         Upgrader.     -   Preferably the hydrotreater charge gas is a bleed gas from other         hydrogenation units in the bitumen Upgrader, allowing for         once-through hydrogen gas configuration, and eliminating the         need for additional make-up compressor facilities or recycle gas         compression or recovery facilities in the bitumen Upgrader.     -   If hydrogen rich waste gas 80 from another hydrotreater is not         available, or in insufficient quantity for once through         operation, or to maximize utilization of available hydrogen in         the waste gas, a recycle gas circuit complete with a compressor         may be employed to recover the hydrogen containing gas from the         hydro-deoxygenation reactor effluent.

The cost of integrating the process of the invention with the virgin oil fractionator processing bitumen is a small fraction of the capital cost of a traditional complete stand-alone hydrogenation unit.

Thus, with the new process flows described herein, which are incorporated into existing bitumen processing flows arranged with the diluted bitumen diluent recovery unit and/or a bitumen vacuum distillation unit, there is no need for any additional process heater, heat exchangers or any additional capacity for waste water treatment, sulfur handling and hydrogen supply.

Now that the arrangement to carry out the process of the invention has been described, persons skilled in the art will readily be able to accommodate known gas oil hydrogenation construction techniques to arrange facilities that carry out the process of the invention.

Pilot plant tests were performed to investigate the reduction of Total Acid Number (TAN) in a blend of bituminous hydrocarbon intermediate streams from a bitumen fractionation unit. The runs were carried out using a ChevronTexaco hydrotreating catalyst (AT-405) operated at a total pressure of 550 psig, over a LHSV range of 2 to 3, and an GHSV range of 500 or 1500 SCF/B, containing 0 to 7% H₂S in the charge gas.

The objectives of the test program were to:

-   -   Demonstrate that low TAN sour virgin distillates can be produced         from high TAN bituminous stocks for production of low TAN sour         synthetic crude blending for sale to the refineries.     -   Measure the TAN reduction kinetics for the blend of Bitumen         derived virgin distillates.     -   Determine the hydrogen consumption of the process.         Experimental Details         Feedstocks

Samples of virgin vacuum LVGO and HVGO distillates derived from Althabasca Bitumen were blended for processing. Detail inspections of the properties of the virgin bitumen derived distillates were performed and the results of the inspections are outlined in Table I below.

Table I—Properties of Bituminous LVGO & HVGO & Blend by Liquid Volume Blend 82 LV % HVGO Feed ID Description HVGO LVGO 18 LV % LVGO API Gravity 14.7 21.5 15.9 Sulfur, Wt % 3.4 2.46 3.23 Viscosity Index −31 16 −12 Viscosity 100° C., cSt 8.96 2.40 7.02 Viscosity 40° C., cSt 138.20 10.36 82.34 TAN, mg KOH/g 4.96 2.91 4.34 Bromine Number, 11.6 gBR/100 g Nitrogen, ppm 1570 478 1330 UV Absorbance Wavelength Gram/Liter Gram/Liter Gram/Liter 226 nm 32.76575 31.43568 32.59277 255 nm 19.55795 11.43124 19.40041 272 nm 14.13724 7.79867 14.06157 305 nm 5.78895 2.77768 5.75063 310 nm 4.65795 1.89056 4.62105 340 nm 1.50828 0.39794 1.50106 348 nm 1.0153 0.24735 1.00717 385 nm 0.17963 0.04768 0.18377 435 nm 0.006 −0.00083 0.01416 450 nm −0.00402 0.00113 0.00875 Simulated distillation TBP (Weight %) ° F. ° F. ° F.  0.5 519 376 466  5 607 460 579 10 636 495 621 30 704 569 702 50 763 617 762 70 817 664 823 90 904 737 906 95 943 782 945 99.5 1044 915 1020

As indicated by the UV Absorbance measurements of the test results given in Table 1, the bitumen compounds are high in ringed unsaturates and the chemistry of these gas oils are very different from the measurements that would typify conventional crude oils. The UV absorbance measurements are typical of absorbance measurements of bitumen derived distillates obtained from the Athabasca tar sands.

Catalyst

A hydrotreating catalyst manufactured by ChevronTexaco and available as their product AT 405 was used to perform pilot tests. The catalyst contains cobalt, molybdenum, and alumina. It was tested as a presulfided {fraction (1/20)}″ diameter cylindrical extrudate, with a nominal Length to diameter ratio of 3 to 4. The catalyst load was 100 cc, 70.4 grams on a dry basis.

Operating Conditions

Table II shows the matrix of operating conditions used in the pilot plant tests and the key results obtained. Note that six sets of conditions were studied. TABLE II Operating Conditions for the Suncor Pilot Plant Tests Run Operating Conditions Run 1 Run 2 Run 3 Run 4 Run 5 Run 6 Reactor Temp, ° F. 485 485 485 525 525 485 LHSV, Hr⁻¹ 2.0 2.0 2.0 2.0 3.0 2.0 Total Pressure, psig 547 551 557 554 544 543 Average H₂ Partial Pressure, psia 310 381 302 305 299 293 Gas/Oil Ratio, SCF/B 500 1500 500 500 500 500 TAN in Product-mg KOH/g 0.65 0.55 0.70 0.35 0.60 0.65 Hydrogen Consumption - SCFB 39 52 41 69 40 40 Changes in [Sulfur] wt % N/C N/C N/C N/C N/C N/C Changes in [Nitrogen] ppmW NE NE NE NE NE NE Changes in [Aromatics] wt % NE NE NE NE NE NE NC = No Change detected NE = No Change Expected due to higher severity requirement than hydro-desulfurization.

The results listed in Table II are presented in graph form in FIG. 4, thus the graph of FIG. 4 presents the same information as Table II and summarizes the results of the six runs graphically.

Based on the results of the pilot testing, the following may be observed:

-   -   At 2.0 LHSV and 500 SCF/B gas/oil ratio, the SOR temperature to         accomplish 90% TAN reduction is 510° F.     -   At 90% TAN reduction, conversion of feed components to products         with boiling points below 650° F. was ˜3 wt %, with a net         chemical H₂ consumption of ˜60 SCF/B.     -   Increasing gas/oil ratio from 500 to 1500 SCF/B increased TAN         reduction activity by ˜15° F.     -   Negligible catalyst deactivation was observed over an 800-hour         period, based on operations at 485° F. and 2.0 LHSV.     -   No detectable changes in Nitrogen, Sulfur and aromatic         concentrations had been observed in the low TAN product samples.

Now that the invention has been described, numerous substitutions, modifications and equivalents will become apparent to those skilled in the art.

The invention is not limited to the preferred embodiments that have been described to illustrate the invention, but rather is defined in the claims appended hereto. 

1. A process for the manufacture of low TAN synthetic crude oil blending components from oil sand bitumen for the production of low TAN synthetic crude oils comprising the steps of: i. providing a supply of hot bitumen distillate; ii. mixing a charge gas containing hydrogen with the bitumen distillate; iii. supplying the mixture under pressure to a hydro-deoxygenating reactor; and iv. recovering the liquid portion of the effluent from the hydro-deoxygenating unit.
 2. The process of claim 1 wherein the supply of hot bitumen distillate is a virgin bitumen distillate or a blend of virgin bitumen distillates produced by atmospheric or vacuum distillation of bitumen.
 3. The process of claim 1 wherein the temperature of the supply of hot bitumen distillate is in the range of 400 to 530 degrees Fahrenheit.
 4. The process of claim 2 wherein the temperature of the supply of hot bitumen distillate is in the range of 400 to 530 degrees Fahrenheit.
 5. The process of claim 1 wherein the charge gas is supplied at a rate that provides hydrogen gas in the range of 250 to 1500 standard cubic feet per barrel of bitumen distillate.
 6. The process of claim 1 wherein the charge gas is obtained from a source of hydrogen gas in a bitumen Upgrader.
 7. The process of claim 6 wherein the source of hydrogen gas is a hydrotreater recycle bleed gas.
 8. The process of claim 7 wherein the hydrogen gas portion of gas from the source of hydrogen gas is supplied at a rate in the range of 250 to 1500 standard cubic feet per barrel of bitumen distillate.
 9. The process according to claim 1, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 10. The process according to claim 6, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 11. The process according to claim 7, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 12. The process according to claim 8, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 13. The process of claim 1 wherein the mixture supplied under pressure to the hydro-deoxygenating unit is supplied at a LHSV in the range of 0.5-5.0.
 14. The process according to claim 2, the hydro-deoxygenating reactor has one or more fixed catalyst beds.
 15. The process of claim 1 wherein the hydro-deoxygenating reactor has a catalyst of nickel-molybdenum or cobalt-molybdenum deposited on an alumina carrier.
 16. The process of claim 14 wherein the hydro-deoxygenating reactor has a catalyst of nickel-molybdenum or cobalt-molybdenum deposited on an alumina carrier.
 17. A process for removing naphthenic acids from bitumen products comprising the steps of: i. supplying bitumen to an atmospheric or vacuum distillation unit to produce a hot gas oil supply; iii. mixing a charge gas containing hydrogen with at least a portion of the hot gas oil supply; iv. supplying the mixture under pressure to a hydro-deoxygenating reactor; and v. recovering the liquid portion of the effluent from the hydro-deoxygenating reactor.
 18. The process of claim 17 wherein the selected portion of the hot gas oil supply in the range of 400 to 530 degrees Fahrenheit.
 19. The process of claim 17 wherein the charge gas is supplied at a rate that provides hydrogen gas in the range of 250 to 1500 standard cubic feet per barrel of bitumen.
 20. The process of claim 17 wherein the charge gas is obtained from a source of hydrogen gas in a bitumen Upgrader.
 21. The process of claim 20 wherein the source of hydrogen gas is a hydrotreater recycle bleed gas.
 22. The process of claim 21 wherein the hydrogen gas portion of gas from the source of hydrogen gas is supplied at a rate in the range of 250 to 1500 standard cubic feet per barrel of bitumen.
 23. The process according to claim 17, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 24. The process according to claim 19, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 25. The process according to claim 20, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 26. The process according to claim 21, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 27. The process according to claim 22, wherein the charge gas is a hydrogen rich gas, containing greater than 50% hydrogen, and 0% to 10% H2S.
 28. The process of claim 17 wherein the mixture supplied under pressure to the hydro-deoxygenating unit is supplied at a LHSV flow rate in the range of 0.5-5.0.
 29. The process of claim 17, wherein the hydro-deoxygenating reactor has one or more fixed catalyst beds.
 30. The process of claim 17 wherein the hydro-deoxygenating reactor has a catalyst of nickel-molybdenum or cobalt-molybdenum deposited on an alumina carrier.
 31. The process of claim 29 wherein the hydro-deoxygenating reactor has a catalyst of nickel-molybdenum or cobalt-molybdenum deposited on an alumina carrier. 